Processes for Isomerizing C8 Aromatic Hydrocarbons Using Serial Reactors
An changeable lead-lag configuration of two isomerization reactors can be used to achieve continuous isomerization operations in an aromatics production complex, even if the isomerization catalyst deactivates over time to require catalyst regeneration and/or replacement. The configuration can be particularly advantageous for two liquid phase isomerization reactors, especially those operated under a high WHSV≥5 hour−1 where the isomerization catalyst can deactivate at a high rate.
This application claims priority to and the benefit of U.S. Ser. No. 62/890,989, filed Aug. 23, 2019, and European Patent Application No. 19209125.4, filed Nov. 14, 2019, the disclosures of which are incorporated herein by their reference.
FIELDThis disclosure relates to isomerization of C8 aromatic hydrocarbons. In particular, this disclosure relates to isomerization of C8 aromatic hydrocarbons under conditions where a majority of the C8 aromatic hydrocarbons are present in liquid phase. This disclosure is useful in, e.g., processes for producing p-xylene.
BACKGROUNDA high purity p-xylene product is typically produced by separating p-xylene from a p-xylene-rich aromatic hydrocarbon mixture comprising p-xylene, o-xylene, m-xylene, and optionally ethylbenzene (“EB”) in a p-xylene separation/recovery system. The p-xylene recovery system can comprise, e.g., a crystallizer and/or an adsorption chromatography separating system known in the prior art. The p-xylene-depleted stream produced from the p-xylene recovery system (the “filtrate” from a crystallizer upon separation of the p-xylene crystals, or the “raffinate” from the adsorption chromatography separating system, collectively “raffinate” in this disclosure) is rich in m-xylene and o-xylene, and contains p-xylene at a concentration typically below its concentration in an equilibrium mixture consisting of m-xylene, o-xylene, and p-xylene. To increase yield of p-xylene, the raffinate stream may be fed into an isomerization unit, where the xylenes undergo isomerization reactions on contacting an isomerization catalyst system to produce an isomerization effluent rich in p-xylene compared to the raffinate. At least a portion of the isomerization effluent, after optional separation and removal of lighter hydrocarbons that may be produced in the isomerization unit, can be recycled to the p-xylene recovery system, forming a “xylenes loop.”
Xylenes isomerization can be carried out under conditions where the C8 aromatic hydrocarbons are substantially in vapor phase in the presence of an isomerization catalyst (vapor-phase isomerization, or “VPI”). References describing VPI processes and/or reactors and/or catalysts include, but are not limited to: U.S. Pat. Nos. 3,651,162, 3,856,872, 3,919,339, 4,098,836, the relevant portions thereof are incorporated herein by reference in their entirety.
Newer generation technology have been developed to allow xylenes isomerization at significantly lower temperature in the presence of an isomerization catalyst, where the C8 aromatic hydrocarbons are substantially in liquid phase (liquid-phase isomerization, or “LPI”). The use of LPI vs. traditional VPI can reduce the number of phase changes (liquid to/from vapor) required to process the C8 aromatic feed. This provides the process with sustainability advantages in the form of significant energy savings. It would be highly advantageous for any p-xylene production plant to deploy a LPI unit, in addition to or in lieu of a VPI unit. For existing p-xylene production facilities lacking a LPI, it would be highly advantageous to add a LPI unit to compliment the VPI unit or replace the VPI unit.
Exemplary LPI processes and catalyst systems useful therefor are described in U.S. Patent Application Publication Nos. 2011/0263918 and 2011/0319688, 2017/0297977, 2016/0257631, and U.S. Pat. No. 9,890,094, the contents of all of which are incorporated herein by reference in their entirety. In the LPI processes described in these references, typically a MFI framework type zeolite (e.g., ZSM-5) is used as the catalyst.
Due to the many advantages of a LPI process and needs to deploy this technology, improvements in this technology, are also needed. It has been found that the LPI catalyst can experience aging over time with various deactivation rate especially at a high weight hourly space velocity (“WHSV”), resulting in reduction of p-xylene concentration in the isomerization effluent from a single reactor over time and the necessary regeneration and/or replacement of the isomerization catalyst when the catalyst has deactivated to a threshold level. It would be highly desirable that in the LPI process, an isomerization effluent comprising p-xylene at a high concentration is consistently produced, and the overall p-xylene production process is not interrupted by the necessary regeneration and/or replacement of the isomerization catalyst.
This disclosure satisfies this and other needs.
SUMMARYIt has been found that by operating at least two LPI reactors (or at least two VPI reactors) connected in series and arranged in alternate lead-lag matter can achieve a consistently high p-xylene concentration in the final isomerization effluent, and maintain operation of at least one LPI reactor (or VPI reactor) during regeneration and/or replacement of the isomerization catalyst in one reactor thereby without interrupting the overall p-xylene production process.
Thus, a first aspect of this disclosure relates to isomerization process, the process comprising one or more of: (I) providing a C8 aromatic hydrocarbon stream; (II) providing a first isomerization reactor comprising a first isomerization catalyst disposed therein, and a second isomerization reactor comprising a second isomerization catalyst disposed therein; (III) feeding the C8 aromatic hydrocarbon stream and optionally molecular hydrogen (H2) into the first isomerization reactor; (IV) contacting the C8 aromatic hydrocarbon stream and the optional molecular hydrogen with the first isomerization catalyst under a first set of isomerization conditions to produce a first isomerization effluent exiting the first isomerization reactor for a first period of time, the first period of time being shorter than the total life cycle of the first isomerization catalyst; (V) obtaining a p-xylene product stream from at least a portion of the first isomerization effluent during the first period of time; (VI) at the end of the first period of time, feeding at least a portion of the first isomerization effluent and optionally additional molecular hydrogen into the second isomerization reactor, wherein at the end of the first period of time, the second isomerization catalyst has a prospective runtime longer than the first isomerization catalyst; (VII) after step (VI), contacting at least a portion of the first isomerization effluent with the second isomerization catalyst under a second set of isomerization conditions in the second isomerization reactor to produce a second isomerization effluent exiting the second isomerization reactor for a second period of time; (VIII) continuing step (IV) during the second period of time; (IX) obtaining a p-xylene product stream from at least a portion of the second isomerization effluent during the second period of time; (X) at the end of the life cycle of the first isomerization catalyst, where the first isomerization catalyst becomes a spent first isomerization catalyst, feeding the C8 aromatic hydrocarbon stream and optionally molecular hydrogen into the second isomerization reactor, and stopping feeding the C8 aromatic hydrocarbon stream into the first isomerization reactor; (XI) after step (X), contacting the C8 aromatic hydrocarbon stream with the second isomerization catalyst under the second set of isomerization conditions to produce a third isomerization effluent exiting the second isomerization reactor for a third period of time; and (XII) obtaining a p-xylene product stream from at least a portion of the third isomerization effluent during the third period of time.
A second aspect of this disclosure relates to C8 aromatic hydrocarbon isomerization process, the process comprising one or more of: (I) providing a C8 aromatic hydrocarbon stream; (II) providing a first isomerization reactor comprising a first isomerization catalyst disposed therein, and a second isomerization reactor comprising a second isomerization catalyst disposed therein; (III) feeding the C8 aromatic hydrocarbon stream and optionally molecular hydrogen (H2) into the first isomerization reactor; (IV) contacting the C8 aromatic hydrocarbon stream and the optional molecular hydrogen with the first isomerization catalyst under a first set of isomerization conditions to produce a first isomerization effluent exiting the first isomerization reactor for a first period of time, the first period of time being shorter than the total life cycle of the first isomerization catalyst; (V) obtaining a p-xylene product stream from at least a portion of the first isomerization effluent during the first period of time; and optionally additional molecular hydrogen at the end of the first period of time, feeding at least a portion of the first isomerization effluent into the second isomerization reactor, wherein at the end of the first period of time, the second isomerization catalyst has an prospective runtime longer than the first isomerization catalyst; (VII) after step (VI), contacting at least a portion of the first isomerization effluent with the second isomerization catalyst under a second set of isomerization conditions in the second isomerization reactor to produce a second isomerization effluent exiting the second isomerization reactor for a second period of time; (VIII) continuing step (IV) during the second period of time; (IX) obtaining a p-xylene product stream from at least a portion of the second isomerization effluent during the second period of time; (X) at the end of the life cycle of the first isomerization catalyst, where the first isomerization catalyst becomes a spent first isomerization catalyst, feeding the C8 aromatic hydrocarbon stream and optionally molecular hydrogen into the second isomerization reactor, and stopping feeding the C8 aromatic hydrocarbon stream into the first isomerization reactor; (XI) after step (X), contacting the C8 aromatic hydrocarbon stream with the second isomerization catalyst under the second set of isomerization conditions to produce a third isomerization effluent exiting the second isomerization reactor for a third period of time; (XII) obtaining a p-xylene product stream from at least a portion of the third isomerization effluent during the third period of time; (XIII) during the third period of time, regenerating the spent first isomerization catalyst in the first isomerization reactor, and/or replacing at least a portion of the spent first isomerization catalyst in the first isomerization reactor with a fresh batch of the first catalyst and/or a regenerated batch of the first catalyst; and (XIV) after step (XIII), designating the second isomerization reactor as the first isomerization reactor, the second isomerization catalyst as the first isomerization catalyst, the second set of isomerization conditions as the first set of isomerization conditions, the first isomerization reactor as the second isomerization reactor, the first isomerization catalyst as the second isomerization catalyst, the first set of isomerization conditions as the second set of isomerization conditions, repeating steps (III), (IV), (V), (VI), (VII), (VIII), (IX), (X), (XI), and (XII), and optionally repeating step (XIII).
Various specific embodiments, versions and examples of the invention will now be described, including preferred embodiments and definitions that are adopted herein for purposes of understanding the claimed invention. While the following detailed description gives specific preferred embodiments, those skilled in the art will appreciate that these embodiments are exemplary only, and that the invention may be practiced in other ways. For purposes of determining infringement, the scope of the invention will refer to any one or more of the appended claims, including their equivalents, and elements or limitations that are equivalent to those that are recited. Any reference to the “invention” may refer to one or more, but not necessarily all, of the inventions defined by the claims.
In this disclosure, a process is described as comprising at least one “step.” It should be understood that each step is an action or operation that may be carried out once or multiple times in the process, in a continuous or discontinuous fashion. Unless specified to the contrary or the context clearly indicates otherwise, multiple steps in a process may be conducted sequentially in the order as they are listed, with or without overlapping with one or more other step, or in any other order, as the case may be. In addition, one or more or even all steps may be conducted simultaneously with regard to the same or different batch of material. For example, in a continuous process, while a first step in a process is being conducted with respect to a raw material just fed into the beginning of the process, a second step may be carried out simultaneously with respect to an intermediate material resulting from treating the raw materials fed into the process at an earlier time in the first step. Preferably, the steps are conducted in the order described.
Unless otherwise indicated, all numbers indicating quantities in this disclosure are to be understood as being modified by the term “about” in all instances. It should also be understood that the precise numerical values used in the specification and claims constitute specific embodiments. Efforts have been made to ensure the accuracy of the data in the examples. However, it should be understood that any measured data inherently contain a certain level of error due to the limitation of the technique and equipment used for making the measurement.
As used herein, the indefinite article “a” or “an” shall mean “at least one” unless specified to the contrary or the context clearly indicates otherwise. Thus, embodiments using “a fractionation column” include embodiments where one, two or more fractionation columns are used, unless specified to the contrary or the context clearly indicates that only one fractionation column is used.
“Consisting essentially of” as used herein means the composition, feed, or effluent comprises a given component at a concentration of at least 60 wt %, preferably at least 70 wt %, more preferably at least 80 wt %, more preferably at least 90 wt %, still more preferably at least 95 wt %, based on the total weight of the composition, feed, or effluent in question.
“Substantially entirely” means at least 95 wt %, preferably ≥98 wt %, preferably ≥99 wt %, preferably ≥99.5 wt %, preferably ≥99.9 wt %. Thus, where C8 aromatic hydrocarbons are present in a reactor substantially entirely in liquid phase, at least 95 wt %, preferably ≥98 wt %, preferably ≥99 wt %, preferably ≥99.5 wt %, preferably ≥99.9 wt %, of such C8 aromatic hydrocarbons are present in the reactor in liquid phase. Where molecular hydrogen is dissolved substantially entirely in an isomerization hydrocarbon feed in liquid phase, at least 95 wt %, preferably ≥98 wt %, preferably ≥99 wt %, preferably ≥99.5 wt %, preferably ≥99.9 wt %, of such molecular hydrogen is dissolved in the isomerization hydrocarbon feed in liquid phase.
The term “hydrocarbon” means (i) any compound consisting of hydrogen and carbon atoms or (ii) any mixture of two or more such compounds in (i). The term “Cn hydrocarbon,” where n is a positive integer, means (i) any hydrocarbon compound comprising carbon atom(s) in its molecule at the total number of n, or (ii) any mixture of two or more such hydrocarbon compounds in (i). Thus, a C2 hydrocarbon can be ethane, ethylene, acetylene, or mixtures of at least two of them at any proportion. A “Cm to Cn hydrocarbon” or “Cm-Cn hydrocarbon,” where m and n are positive integers and m<n, means any of Cm, Cm+1, Cm+2, . . . , Cn−1, Cn hydrocarbons, or any mixtures of two or more thereof. Thus, a “C2 to C3 hydrocarbon” or “C2-C3 hydrocarbon” can be any of ethane, ethylene, acetylene, propane, propene, propyne, propadiene, cyclopropane, and any mixtures of two or more thereof at any proportion between and among the components. A “saturated C2-C3 hydrocarbon” can be ethane, propane, cyclopropane, or any mixture thereof of two or more thereof at any proportion. A “Cn+ hydrocarbon” means (i) any hydrocarbon compound comprising carbon atom(s) in its molecule at the total number of at least n, or (ii) any mixture of two or more such hydrocarbon compounds in (i). A “Cn− hydrocarbon” means (i) any hydrocarbon compound comprising carbon atoms in its molecule at the total number of at most n, or (ii) any mixture of two or more such hydrocarbon compounds in (i). A “Cm hydrocarbon stream” means a hydrocarbon stream consisting essentially of Cm hydrocarbon(s). A “Cm-Cn hydrocarbon stream” means a hydrocarbon stream consisting essentially of Cm-Cn hydrocarbon(s).
“Light hydrocarbon” in this disclosure means any C5− hydrocarbon.
“Liquid-phase isomerization” and “LPI” interchangeably mean a C8 aromatic hydrocarbon isomerization process in an isomerization reactor in the presence of an isomerization catalyst whereby the xylenes isomerize under isomerization conditions such that the C8 aromatic hydrocarbons present in the isomerization reactor are substantially in liquid phase. “Substantially in liquid phase” means ≥80 wt %, preferably ≥85 wt %, preferably ≥90 wt %, preferably ≥95 wt %, preferably ≥99 wt %, preferably the entirety, are in liquid phase. Such isomerization conditions are called liquid-phase isomerization conditions.
“Vapor-phase isomerization” and “VPI” interchangeably mean a C8 aromatic hydrocarbon isomerization process in an isomerization reactor in the presence of an isomerization catalyst whereby the xylenes isomerize under isomerization conditions such that the xylenes present in the isomerization reactor are substantially in vapor phase. “Substantially in vapor phase” means ≥90 wt %, preferably ≥95 wt %, preferably ≥99 wt %, preferably the entirety, is in vapor phase. Such isomerization conditions are called vapor-phase isomerization conditions.
As used herein, “wt %” means percentage by weight, “vol %” means percentage by volume, “mol %” means percentage by mole, “ppm” means parts per million, and “ppm wt” and “wppm” are used interchangeably to mean parts per million on a weight basis. All concentrations herein are expressed on the basis of the total amount of the composition in question. All ranges expressed herein should include both end points as two specific embodiments unless specified or indicated to the contrary.
Nomenclature of elements and groups thereof used herein are pursuant to the Periodic Table used by the International Union of Pure and Applied Chemistry after 1988. An example of the Periodic Table is shown in the inner page of the front cover of Advanced Inorganic Chemistry, 6th Edition, by F. Albert Cotton et al. (John Wiley & Sons, Inc., 1999).
The processes of this disclosure can be used for making a p-xylene product comprising p-xylene at a high concentration, and/or an o-xylene product comprising o-xylene at a high concentration, and/or an m-xylene product comprising m-xylene at a high concentration. Given the higher economic value of p-xylene and o-xylene over m-xylene, the processes of this disclosure are preferably used for making p-xylene and/or o-xylene products, more preferably a p-xylene product. In typical C8 aromatic hydrocarbon isomerization processes, an isomerization hydrocarbon feed comprising the xylenes at non-equilibrium concentrations are contacted with an isomerization catalyst in an isomerization reactor under isomerization conditions for a period of time to effect the conversion of the xylenes to produce an isomerization effluent exiting the isomerization reactor comprising the xylenes at concentrations closer to their equilibrium concentrations. Thus, for the purpose of producing a p-xylene product, the isomerization hydrocarbon feed fed into the isomerization reactor typically comprises (i) p-xylene at a concentration based on the total quantity of xylenes in the isomerization hydrocarbon feed significantly lower than its equilibrium concentration in a xylenes mixture, and (ii) o-xylene and m-xylene at a combined concentration based on the total quantity of xylenes in the isomerization hydrocarbon feed significantly higher than their combined equilibrium concentrations. On contacting the isomerization catalyst in the isomerization reactor under the isomerization conditions, a portion of the m-xylene and/or o-xylene is converted into p-xylene to produce an isomerization effluent comprising p-xylene at a higher concentration thereof based on the total quantity of the xylenes in the effluent. It is highly desired that the p-xylene concentration based on the total quantity of xylenes in the isomerization effluent is close to its equilibrium concentration in a xylenes mixture. The increased amount of p-xylene in the effluent can be recovered to produce a p-xylene product. A “xylenes mixture” means a mixture consisting of p-xylene, m-xylene, and o-xylene. An “equilibrium xylenes mixture” means a xylenes mixture comprising the three xylene isomers at thermodynamic equilibrium.
The isomerization catalyst used in a VPI or LPI process can deactivate overtime due to, e.g., coke formation on the catalyst. After the catalyst deactivates to a certain degree, the isomerization catalyst may require regeneration and/or replacement to ensure the isomerization effluent comprises p-xylene at a desired concentration. If a single isomerization reactor is utilized in the aromatics production complex, the interruption of the operation of the isomerization reactor can interrupt the operation of the whole complex, which is highly undesirable. It would be highly desirable to operate an isomerization process at a WHSV of the isomerization hydrocarbon feed of, e.g., ≥5 hour−1, ≥10 hour−1, ≥15 hour−1, or even ≥20 hour−1, because a high WHSV enables a small isomerization reactor size. However, a high WHSV isomerization process, especially a LPI process, can cause a fast deactivation of the isomerization catalyst, necessitating relatively short catalyst cycle of the catalyst, and relatively frequent catalyst regeneration and/or change-out. It would be highly desirable that the shut-down of an isomerization reactor, such as a LPI reactor, would have minimal impact on the continuous operation of the aromatic hydrocarbon production complex including that isomerization reactor.
The processes of this disclosure, by utilizing at least two isomerization reactors, with two of them operated in lead-lag fashion that can be alternated, have the capability to provide a continuous isomerization operation permitting periodic shut-down of one isomerization reactor and isomerization catalyst regeneration and/or replacement. The processes are particularly suitable for LPI reactors, more particularly those running at high WHSV experiencing appreciable isomerization catalyst deactivation. The processes of this disclosure are advantageously continuous. Thus, multiple steps as described in the various embodiments of the processes may be carried out simultaneously.
Thus, various embodiments of the isomerization process of this disclosure comprise: (I) providing an isomerization hydrocarbon feed comprising C8 aromatic hydrocarbons; (II) providing a first isomerization reactor comprising a first isomerization catalyst disposed therein, and a second isomerization reactor comprising a second isomerization catalyst disposed therein; (III) feeding the isomerization hydrocarbon feed and optionally molecular hydrogen (H2) into the first isomerization reactor; (IV) contacting the isomerization hydrocarbon feed and the optional molecular hydrogen with the first isomerization catalyst under a first set of isomerization conditions to produce a first isomerization effluent exiting the first isomerization reactor for a first period of time, the first period of time being shorter than the total life cycle of the first isomerization catalyst; (V) obtaining a p-xylene product stream from at least a portion of the first isomerization effluent during the first period of time; (VI) at the end of the first period of time, feeding at least a portion of the first isomerization effluent and optionally additional molecular hydrogen into the second isomerization reactor, wherein at the end of the first period of time, the second isomerization catalyst has a prospective runtime longer than the first isomerization catalyst; (VII) after step (VI), contacting at least a portion of the first isomerization effluent with the second isomerization catalyst under a second set of isomerization conditions in the second isomerization reactor to produce a second isomerization effluent exiting the second isomerization reactor for a second period of time; (VIII) continuing step (IV) during the second period of time; (IX) obtaining a p-xylene product stream from at least a portion of the second isomerization effluent during the second period of time; (X) at the end of the life cycle of the first isomerization catalyst, where the first isomerization catalyst becomes a spent first isomerization catalyst, feeding the isomerization hydrocarbon feed and optionally molecular hydrogen into the second isomerization reactor, and stopping feeding the isomerization hydrocarbon feed into the first isomerization reactor; (XI) after step (X), contacting the isomerization hydrocarbon feed with the second isomerization catalyst under the second set of isomerization conditions to produce a third isomerization effluent exiting the second isomerization reactor for a third period of time; and (XII) obtaining a p-xylene product stream from at least a portion of the third isomerization effluent during the third period of time. In these embodiments, the first isomerization reactor can be considered as a lead reactor, and the second isomerization reactor a lag reactor.
In certain embodiments, the isomerization process of this disclosure can further comprise: (XIII) during the third period of time, regenerating the spent first isomerization catalyst in the first isomerization reactor, and/or replacing at least a portion of the spent first isomerization catalyst in the first isomerization reactor with a fresh batch of the first catalyst and/or a regenerated batch of the first catalyst. Regeneration of the first isomerization catalyst can be performed in situ in the first isomerization reactor by, e.g., exposing the spent first isomerization catalyst to a stream (e.g., a vapor stream, a liquid stream, or a liquid/vapor mixture stream) of molecular hydrogen-containing-rich fluid (e.g., a high-purity molecular hydrogen vapor stream, a molecular hydrogen/inert gas vapor mixture stream, a molecular hydrogen/hydrocarbon mixture stream, and the like) at an elevated temperature. Additionally or alternatively, the spent first isomerization catalyst or a portion thereof may be exposed to an oxygen-containing atmosphere in-situ in the first isomerization reactor, and/or ex-situ outside of the first isomerization reactor to at least partly effect the regeneration of the spent first isomerization catalyst. If regenerated ex situ, the regenerated first isomerization catalyst can be charged into the first isomerization reactor, optionally together with a quantity of the first isomerization catalyst.
In certain embodiments of the isomerization process of this disclosure the third period of time ends before the end of the life cycle of the second isomerization catalyst. In certain embodiments, the process further comprises: (XIV) after step (XIII), designating the second isomerization reactor as the first isomerization reactor, the second isomerization catalyst as the first isomerization catalyst, the second set of isomerization conditions as the first set of isomerization conditions, the first isomerization reactor as the second isomerization reactor, the first isomerization catalyst as the second isomerization catalyst, the first set of isomerization conditions as the second set of isomerization conditions, and subsequently repeating steps (III), (IV), (V), (VI), (VII), (VIII), (IX), (X), (XI), and (XII). Thus, the original lag isomerization reactor becomes a lead reactor, and the original lead isomerization reactor becomes a lag reactor. The positions of the two reactors are alternated. In certain embodiments, step (XIV) further comprises repeating step (XIII). At the end of the catalyst life cycle of the new lead reactor, it may be shut down to allow for the regeneration and/or replacement of the spent isomerization catalyst therein. By repeating these steps, the two isomerization reactors can alter the lead/lag positions many times, while allowing at any given time at least one of them in operation with a desirable isomerization catalyst activity. In certain embodiments, the process further comprises (XV) stopping operating the first isomerization reactor and/or the second isomerization reactor at a major turn-around point of time of the respective isomerization reactor(s). Thus, the two isomerization reactors may be both taken down at a time when the overall aromatics product complex is due for a major overhaul, after the two isomerization reactors have alternated the lead-lag arrangement multiple times.
In certain embodiments of the process of this disclosure, the first period of time is at least half of the life cycle of the first isomerization catalyst. Thus, for example, the first period of time can be ≥60%, ≥70%, ≥75%, ≥80%, ≥85%, ≥90% of the life cycle of the first isomerization catalyst. To the extent during the first period of time, the performance of the first isomerization catalyst remains acceptable by, e.g., producing an isomerization effluent at an acceptable catalyst activity, it may be advantageous to allow the first period of time to extend as long as practical.
In certain embodiments of the isomerization process of this disclosure, at the end of the first period of time, in step (IV), the activity of the first isomerization catalyst is no greater than a threshold percentage of the average activity of the first isomerization catalyst during the first period of time. The activity of the first isomerization catalyst can be expected to decrease over time during the first period of time. When the activity decreases to a threshold level of the average of the first period of time, the first isomerization effluent may comprise p-xylene at an undesirably low concentration. At that point, at least a portion of the first isomerization effluent is fed into the second isomerization reactor in step (VI). Inside the second isomerization reactor, additional isomerization of the C8 aromatic hydrocarbons in the first isomerization effluent can take place to produce more p-xylene product in the second isomerization reactor, resulting in a high concentration of p-xylene in the second isomerization effluent. The threshold percentage can be selected at or between any of, e.g., 50%, 60%, 70%, 80%, 85%, 90%, 95%, and the like.
In certain embodiments, the first isomerization catalyst and the second isomerization catalyst have substantially the same composition and therefore similar performance under the same isomerization conditions. In other embodiments, the first isomerization catalyst and the second isomerization catalyst can have very different compositions.
The isomerization process of this disclosure is particularly advantageous where the first and/or the second isomerization conditions include a high WHSV of, e.g., ≥5 hour−1, ≥7.5 hour−1, ≥10 hour−1, ≥12.5 hour−1, ≥15 hour−1, ≥17.5 hour−1, and ≥20 hour−1, where the isomerization catalyst(s) can experience a pronounced deactivation rate at such high throughput.
In certain embodiments of the process of this disclosure, the first set of isomerization conditions are such that vapor-phase isomerization is carried out in the first isomerization reactor; and the second set of isomerization conditions are such that vapor-phase isomerization is carried out in the second isomerization reactor. In such embodiments, both the first and second isomerization reactors are VPI reactors operated under VPI conditions. In such embodiments, the first and second isomerization reactors may alternate as lead or lag reactors during many cycles of operations of both, enabling an overall continuous operation of VPI processes permitting catalyst regeneration and/or change-out without interrupting the operation of other components of the aromatics production complex including the two VPI reactors. In certain embodiments, the two VPI reactors have similar capacities and/or are designed to be operated under similar isomerization conditions. In other embodiments, the two VPI reactors can have different capacities and/or are designed to be operated under different isomerization conditions.
In certain embodiments of this disclosure, the first set of isomerization conditions are such that liquid-phase isomerization is carried out in the first isomerization reactor; and the second set of isomerization conditions are such that liquid-phase isomerization is carried out in the second isomerization reactor. In such embodiments, both the first and second isomerization reactors are LPI reactors operated under the same or different LPI conditions. In such embodiments, the first and second isomerization reactors may alternate as lead or lag reactors during many cycles of operations of both, enabling an overall continuous operation of LPI processes permitting catalyst regeneration and/or change-out without interrupting the operation of other components of the aromatics production complex including the two LPI reactors. In certain embodiments, the two LPI reactors have similar capacities and/or are designed to be operated under similar isomerization conditions. In other embodiments, the two LPI reactors can have different capacities and/or are designed to be operated under different isomerization conditions. In certain embodiments, the first and second isomerization conditions in the two LPI reactors comprise at least one of the following: a temperature in a range from 200 to 300° C. (preferably 250 to 300° C.), a weight hourly space velocity in a range from 1.5 to 20 hour−1 (preferably 2.5 to 20 hour−1, preferably 5.0 to 20 hour−1, preferably 10.0 to 20 hour−1), and a gauge pressure in a range from 0 to 3500 kilopascal (preferably from 690 to 3500 kilopascal, preferably from 1380 to 3500 kilopascal). In such LPI processes, molecular hydrogen may or may not be fed into one or both of the isomerization reactors. In certain embodiments, in step (III), molecular hydrogen is fed into the first isomerization reactor at a first hydrogen feeding rate; and in step (X), molecular hydrogen is fed into the second isomerization reactor at a second hydrogen feeding rate. In certain embodiments, the first hydrogen feeding rate and the second hydrogen feeding rate are in a range below 100 ppm by weight, based on the total weight of the isomerization hydrocarbon feed. In certain embodiments, the first set of isomerization conditions and the second set of isomerization conditions comprise a WHSV of less than 5 hour−1. In certain embodiments, the first hydrogen feeding rate and the second hydrogen feeding rate are in a range from 100 to 5000 by weight, based on the total weight of the isomerization hydrocarbon feed. In certain embodiments, the first set of isomerization conditions and the second set of isomerization conditions comprise a WHSV from 5 to 25 hour−1 (e.g, 10 to 25 hour−1).
In certain embodiments of the isomerization process of this disclosure, the first and second isomerization catalysts comprise a zeolite having a 10 to 12-member ring structure therein.
In certain embodiments of the isomerization process of this disclosure, the first and second isomerization catalysts are free of any precious metal.
Certain preferred embodiments of the isomerization process of this disclosure can comprise: (I) providing an isomerization hydrocarbon feed comprising C8 aromatic hydrocarbons; (II) providing a first isomerization reactor comprising a first isomerization catalyst disposed therein, and a second isomerization reactor comprising a second isomerization catalyst disposed therein; (III) feeding the isomerization hydrocarbon feed and optionally molecular hydrogen (H2) into the first isomerization reactor; (IV) contacting the isomerization hydrocarbon feed and the optional molecular hydrogen with the first isomerization catalyst under a first set of isomerization conditions to produce a first isomerization effluent exiting the first isomerization reactor for a first period of time, the first period of time being shorter than the total life cycle of the first isomerization catalyst; (V) obtaining a p-xylene product stream from at least a portion of the first isomerization effluent during the first period of time; (VI) at the end of the first period of time, feeding at least a portion of the first isomerization effluent and optionally additional molecular hydrogen into the second isomerization reactor, wherein at the end of the first period of time, the second isomerization catalyst has an prospective runtime longer than the first isomerization catalyst; (VII) after step (VI), contacting at least a portion of the first isomerization effluent with the second isomerization catalyst under a second set of isomerization conditions in the second isomerization reactor to produce a second isomerization effluent exiting the second isomerization reactor for a second period of time; (VIII) continuing step (IV) during the second period of time; (IX) obtaining a p-xylene product stream from at least a portion of the second isomerization effluent during the second period of time; (X) at the end of the life cycle of the first isomerization catalyst, where the first isomerization catalyst becomes a spent first isomerization catalyst, feeding the isomerization hydrocarbon feed and optionally molecular hydrogen into the second isomerization reactor, and stopping feeding the isomerization hydrocarbon feed into the first isomerization reactor; (XI) after step (X), contacting the isomerization hydrocarbon feed with the second isomerization catalyst under the second set of isomerization conditions to produce a third isomerization effluent exiting the second isomerization reactor for a third period of time; (XII) obtaining a p-xylene product stream from at least a portion of the third isomerization effluent during the third period of time; (XIII) during the third period of time, regenerating the spent first isomerization catalyst in the first isomerization reactor, and/or replacing at least a portion of the spent first isomerization catalyst in the first isomerization reactor with a fresh batch of the first catalyst and/or a regenerated batch of the first catalyst; and (XIV) after step (XIII), designating the second isomerization reactor as the first isomerization reactor, the second isomerization catalyst as the first isomerization catalyst, the second set of isomerization conditions as the first set of isomerization conditions, the first isomerization reactor as the second isomerization reactor, the first isomerization catalyst as the second isomerization catalyst, the first set of isomerization conditions as the second set of isomerization conditions, repeating steps (III), (IV), (V), (VI), (VII), (VIII), (IX), (X), (XI), and (XII), and optionally repeating step (XIII).
In certain embodiments of the processes of this disclosure, one can increase the reaction temperature and/or the weight hourly space velocity without substantially decreasing the p-Xylene/Xylenes weight ratio in the isomerization effluent. “Substantially decreasing the same p-Xylene/Xylenes weight ratio” means decreasing the p-Xylene/Xylenes weight ratio in the isomerization effluent by ≥10% based on the p-Xylene/Xylenes weight ratio before increasing the reaction temperature and/or the weight hourly space velocity. “p-Xylene/Xylenes” weight ratio means the weight ratio of p-xylene to all xylenes present in a mixture or stream.
In certain other embodiments of the processes of this disclosure, one can increase the reaction temperature and/or the weight hourly space velocity without substantially increasing the xylenes loss in the process. “Substantially increasing the xylenes loss” means increasing the xylenes loss by ≥10% based on the xylenes loss before increasing the reaction temperature and/or the weight hourly space velocity. “Xylenes loss” means the concentration of all xylenes present in an effluent stream minus the concentration of all xylenes present in the feed, in weight percentage.
In certain other embodiments of the processes of this disclosure, one can feed the molecular hydrogen at the beginning phase of a catalyst cycle at a first rate, and then subsequently during normal operation of the isomerization reactor at a second rate, where the first rate is lower than the second rate. The beginning phase of a catalyst cycle means the period of a catalyst cycle when the catalyst demonstrates an exceedingly high activity which can result in side-reactions and over-production of byproducts. For example, in the beginning phase, the molecular hydrogen can be fed into reactor at a rate from 0 to 500, or from 0 to 300, or from 0 to 200, or from 0 to 150 ppm. At a low first feeding rate of the molecular hydrogen, the isomerization catalyst is allowed to deactivate (“de-edge”) at a relatively high rate to reduce the side-reaction and the production of byproducts. At the end of the beginning phase, where de-edging of the isomerization catalyst is complete, the molecular hydrogen feeding rate can be increased to a higher level to reduce the deactivation of the isomerization catalyst to a desired level.
The Isomerization Hydrocarbon Feed Comprising C8 Aromatic Hydrocarbon in the Processes of this Disclosure
The isomerization hydrocarbon feed comprising C8 aromatic hydrocarbon to the isomerization reactor(s) can comprise one of the more of the xylenes at various concentrations. For example, the isomerization hydrocarbon feed can comprise xylenes at a total concentration from c(xylenes)1 to c(xylenes)2 wt %, based on the total weight of the isomerization hydrocarbon feed, where c(xylenes)1 and c(xylenes)2 can be, independently, 30, 40, 50, 55, 60, 65, 70, 75, 80, 85, 90, 91, 92, 93, 94, 95, 96, 97, 98, 99, or even 100, as long as c(xylenes)1<c(xylenes)2. Preferably c(xylenes)1=70. More preferably c(xylenes)1=80. Preferably the isomerization hydrocarbon feed consists essentially of the xylenes.
The isomerization hydrocarbon feed can comprise ethylbenzene at various concentrations, e.g., from c(EB)1 to c(EB)2 wt %, based on the total weight of the isomerization hydrocarbon feed, where c(EB)1 and c(EB)2 can be, independently, 1, 2, 3, 4, 5, 6, 7, 8, 9, 10, 11, 12, 13, 14, 15, 16, 17, 18, 19, 20, 21, 22, 23, 24, 25, 26, 27, 28, 29, 30, as long as c(EB)1<c(EB)2. Preferably C(EB)2=20. Preferably C(EB)2=10. Preferably C(EB)2=5.
For the purpose of producing a p-xylene product, it is highly desirable that the isomerization hydrocarbon feed comprises p-xylene at a concentration based on the total quantity of xylenes in the isomerization hydrocarbon feed lower than the concentration thereof in an equilibrium xylenes mixture. Thus p-xylene in the isomerization hydrocarbon feed can have a concentration c(pX) based on the total quantity of the xylenes in the isomerization hydrocarbon feed ranging from c(pX)1 to c(pX)2 wt %, where c(pX)1 and c(pX)2 can be, independently, e.g., 0, 0.5, 1, 2, 3, 4, 5, 6, 7, 8, 9, 10, 11, 12, 13, 14, 15, 16, 17, 18, 19, 20, 21, 22, 23, 24, 25, 26, 27, 28, 29, 30, as long as c(pX)1<c(pX)2. Preferably c(pX)2=20. Preferably c(pX)2=15. Preferably c(pX)2=12. Preferably c(pX)2=10. Preferably c(pX)2=8. Preferably c(pX)2=6. Preferably c(pX)2=5. Preferably c(pX)2=4. Preferably c(pX)2=2. Preferably c(pX)2=1. m-Xylene in the isomerization hydrocarbon feed can have a concentration c(mX) based on the total quantity of the xylenes in the isomerization hydrocarbon feed ranging from c(mX)1 to c(mX)2 wt %, where c(mX)1 and c(mX)2 can be, independently, e.g., 10, 15, 20, 25, 30, 35, 40, 45, 50, 55, 60, 65, 70, 75, 80, 85, 90, 95, as long as c(mX)1<c(mX)2. Preferably c(mX)1=30 and c(mX)2=80. Preferably c(mX)1=40, and c(mX)2=80. o-Xylene in the isomerization hydrocarbon feed can have a concentration c(oX) based on the total quantity of the xylenes in the isomerization hydrocarbon feed ranging from c(oX)1 to c(oX)2 wt %, where c(oX)1 and c(oX)2 can be, independently, e.g., 10, 15, 20, 25, 30, 35, 40, 45, 50, 55, 60, 65, 70, 75, 80, 85, 90, 95, as long as c(oX)1<c(oX)2. Preferably c(oX)1=10 and c(oX)2=80. Preferably c(oX)1=10, and c(oX)2=60. Preferably c(oX)1=10, and c(oX)2=50. Preferably c(oX)1=15, and c(oX)2=30. Such feed can be a raffinate stream exiting a p-xylene recovery sub-system, e.g., one using adsorption chromatography-based separation technology.
The isomerization hydrocarbon feed may comprise benzene, toluene, and C9+ hydrocarbons, but desirably at low quantities. The isomerization hydrocarbon feed may comprise benzene and toluene combined in a range from c(BT)1 to c(BT)2 wt %, based on the total weight of the isomerization hydrocarbon feed, where c(BT)1 and c(BT)2 can be, independently, e.g., 0.01, 0.1, 1.0, 2.0, 3.0, 5.0, 8.0, 10.0, 15.0, 20.0, as long as c(BT)1<c(BT)2. Preferably c(BT)2=10.0. Preferably c(BT)2=5.0. Preferably c(BT)2=3.0. Toluene can be the primary component between benzene and toluene. For example, the isomerization hydrocarbon feed can be substantially free of benzene. The isomerization hydrocarbon feed may comprise C9+ hydrocarbons, in total, in a range from c(C9+)1 to c(C9+)2 wt %, based on the total weight of the isomerization hydrocarbon feed, where c(C9+)1 and c(C9+)2 can be, independently, e.g., 0.01, 0.1, 1.0, 5.0, 10.0, 20.0, as long as c(C9+)1<c(C9+)2.
In various preferred embodiments of the processes of this disclosure, at least 80 wt %, preferably ≥85 wt %, preferably ≥90 wt %, preferably 95 wt %, preferably ≥98 wt %, preferably ≥99 wt %, preferably approximately 100 wt %, of the isomerization hydrocarbon feed comprising the C8 aromatic hydrocarbons is in liquid phase at the inlet to the first isomerization reactor and/or the second isomerization reactor. In such embodiments, the isomerization hydrocarbon feed can have an inlet temperature in the range from T1 to T2° C., where T1 and T2 can be, independently, e.g., 200, 210, 220, 230, 240, 250, 255, 260, 265, 270, 275, 280, 285, 290, 295, 300, as long as T1<T2. The relatively low inlet temperature of the isomerization hydrocarbon feed, in combination with other isomerization conditions such as pressure described below, can enables an LPI process inside the first and second isomerization reactors.
The Molecular Hydrogen Co-Feed
Molecular hydrogen may or may not be fed into the first and/or second isomerization reactors of the isomerization process of this disclosure. The molecular hydrogen, if co-fed into the first isomerization reactor and/or the second isomerization reactor, can be injected into the respective isomerization reactor via an inlet as a pressurized gas. Additionally or alternatively, the molecular hydrogen or a portion thereof can be fed into a feeding line, a vessel, or a storage tank of the isomerization hydrocarbon feed comprising the C8 aromatic hydrocarbons, where it forms a mixture with the isomerization hydrocarbon feed, which is then supplied into the isomerization reactor.
VPI processes can be preformed in the presence of molecular hydrogen co-fed into the first and/or second isomerization reactors along with the isomerization hydrocarbon feed. VPI processes and reactors known in the art may be used. References describing VPI processes and/or reactors and/or catalysts include, but are not limited to: U.S. Pat. Nos. 3,651,162, 3,856,872, 3,919,339, 4,098,836, 7,247,762, and 7,271,118, the relevant portions thereof are incorporated herein by reference in their entirety.
Certain LPI processes in the prior art were conducted in the absence of co-feeding molecular hydrogen into the isomerization reactor. We have found that in the absence of co-feeding any molecular hydrogen into the reactor, the isomerization catalyst can deactivate overtime at a relatively fast pace, especially at a high WHSV≥5 hour−1. In cases where molecular hydrogen is supplied into the isomerization reactor at a low feeding rate, e.g., at ≤10 ppm by weight, based on the total weight of the aromatic hydrocarbon feed into the reactor, the isomerization catalyst may deactivate at a pace which, compared to no co-feeding of molecular hydrogen at all, is lower, but nonetheless can be appreciable. Isomerization catalyst deactivation at an appreciable pace can be observed at low feeding rate of molecular hydrogen or in the absence of co-feeding molecular hydrogen even in cases where the isomerization hydrocarbon feed is relatively easy to isomerize, e.g., where the isomerization feed comprises p-xylene at a high concentration (e.g., ≥5 wt %, ≥8 wt %, ≥10 wt %, based on the total weight of the xylenes) and/or the isomerization feed comprises ethylbenzene at a low concentration (e.g., ≤8 wt %, ≤6 wt %, ≤5 wt %, ≤4 wt %, ≤2 wt %).
It was surprisingly found that when molecular hydrogen is co-fed at a high feeding rate of ≥100 ppm by weight into an LPI reactor, based on the total weight of the isomerization hydrocarbon feed, the deactivation rate of the LPI catalyst can be reduced significantly compared to both (i) no co-feeding of molecular hydrogen at all and (ii) co-feeding molecular hydrogen at a low rate, e.g., at ≤10 ppm by weight. In a totally unexpected manner, the low deactivation rate with co-feeding of molecular hydrogen at ≥100 ppm was observed even at high WHSV≥5 hour−1, ≥7.5 hour−1, ≥10 hour−1, ≥12.5 hour−1, ≥15 hour−1, ≥17.5 hour−1, and even ≥20 hour−1.
It is highly desired that a portion, preferably a majority (e.g., ≥50%, ≥60%, ≥70%, ≥80%, ≥90%, ≥95%, ≥98%), more preferably substantially the entirety (≥99%), of the co-fed molecular hydrogen is dissolved in the liquid phase in the LPI reactor(s). To achieve a higher concentration of dissolved molecular hydrogen in the liquid phase of the isomerization hydrocarbon feed, a higher pressure can be applied. To maintain a substantial portion of the molecular hydrogen dissolved in the liquid phase in an LPI reactor, it is highly desired that the feeding rate of molecular hydrogen into the isomerization reactor be no higher than 5000 ppm by weight, based on the total weight of the isomerization hydrocarbon feed. Approximate solubility of molecular hydrogen in liquid-phase C8 aromatic hydrocarbons in ppm by weight, based on the total weight of the C8 aromatic hydrocarbons, at 25° C. at various pressures is given in TABLE I below:
Thus, the molecular hydrogen can be fed into an LPI reactor at a feeding rate of r(H2)1 to r(H2)2 ppm by weight, based on the total weight of the isomerization hydrocarbon feed, where r(H2)1 and r(H2)2 can be, independently, e.g., 100, 150, 200, 250, 300, 350, 400, 450, 500, 550, 600, 650, 700, 750, 800, 850, 950, 1000, 1500, 2000, 2500, 3000, 3500, 4000, 4500, 5000, as long as r(H2)1<r(H2)2. Preferably r(H2)2=3000. Preferably r(H2)2=2000. Preferably r(H2)2=1000. Preferably r(H2)2=800. Preferably r(H2)2=600. Preferably r(H2)2=500.
The molecular hydrogen, if premixed with the isomerization hydrocarbon feed, will be fed into the isomerization reactor at an inlet temperature as discussed above. If fed separately from the isomerization hydrocarbon feed, it may be fed into the isomerization reactor as a stream gas at an inlet temperature preferably in proximity to the inlet temperature of the isomerization hydrocarbon feed, at a pressure sufficient to enable its dissolution in the liquid phase in the isomerization reactor to effect LPI in the isomerization reactor.
In particularly advantageous embodiments of the processes of this disclosure, the first isomerization reactor and the second isomerization reactors are both LPI reactors, and molecular hydrogen is co-fed into at least one, preferably both, of the LPI reactors at a significant quantity of ≥100 ppm by weight, based on the total weight of the hydrocarbon feed fed into the respective LPI reactor. It has been surprisingly found that the presence of such significant amount of molecular hydrogen can significantly reduce the deactivation rate of the LPI catalyst, even at a high WHSV≥5.0, ≥7.5, ≥10, ≥12.5, ≥15, ≥17.5, and even ≥20 hour−1.
Various embodiments of the processes of this disclosure comprise introducing a feed comprising C8 hydrocarbons and molecular hydrogen into a reactor having a xylene isomerization catalyst disposed therein, wherein the isomerization hydrocarbon feed comprises 100 ppm to 5000 ppm by weight of the molecular hydrogen based on the total weight of the isomerization hydrocarbon feed, and reacting the isomerization hydrocarbon feed substantially in liquid phase in the presence of the xylene isomerization catalyst to produce an isomerization effluent having an increased concentration of para-xylene relative to a concentration of para-xylene in the isomerization hydrocarbon feed, wherein the isomerization hydrocarbon feed contacts the xylene isomerization catalyst under a gauge pressure of about 1,700 kPa to about 3,500 kPa at a weight hour space velocity of about 5 hour−1 to about 20 hour−1.
Preferred LPI conditions in the first and/or second LPI reactors can include a reaction gauge pressure in the isomerization reactor ranging from p1 to p2 kPa, where p1 and p2 can be, independently, e.g., 1700, 1800, 1900, 2000, 2100, 2200, 2300, 2400, 2500, 2600, 2700, 2800, 2900, 3000, 3100, 3200, 3300, 3400, 3500, as long as p1<p2. Preferably p2=3000. Preferably p2=2500. The higher the reaction pressure, the larger the quantity of molecular hydrogen can be dissolved in the liquid phase of the hydrocarbons present in the reactor, as indicated above.
Preferred LPI conditions in the first and/or second LPI reactors can include a reaction temperature in the isomerization reactor ranging from T1 to T2° C., where T1 and T2 can be, independently, e.g., 200, 210, 220, 230, 240, 250, 255, 260, 265, 270, 275, 280, 285, 290, 295, 300, as long as T1<T2. The relatively low reaction temperature of the LPI processes can be particularly energy efficient because it requires less energy to heat the isomerization hydrocarbon feed, and it does not require condensing a large quantity of high-temperature vapor-phase isomerization effluent into liquid for downstream processing.
Preferred LPI conditions in the first and/or second LPI reactors can particularly advantageously include a high WHSV ranging from w1 to w2 hour−1, where w1 and w2 can be, e.g., 5.0, 5.5, 6.0, 6.5, 7.0, 7.5, 8.0, 8.5, 9.0, 9.5, 10, 11, 12, 12.5, 13, 14, 15, 16, 17, 17.5, 18, 19, 20, as long as w1<w2. As indicated above and demonstrated in the comparative examples below, LPI processes without co-feeding molecular hydrogen or with co-feeding molecular hydrogen at a low feeding rate can experience catalyst deactivation at a high rate, even at a WHSV of 4 hour−1. Such high WHSV of the processes of this disclosure enables compact LPI reactor design for a new LPI with a given capacity, increased capacity for an existing LPI reactor with a given catalyst load, and reduced catalyst consumption for the production of a given quantity of p-xylene product.
U.S. Pat. Nos. 6,180,550, 6,448,459, 6,872,866, 7,244,409, 7,371,913, 7,495,137, 7,592,499, 8,221,707, 8,273,934, and 8,697,929 describe LPI processes and/or catalysts suitable for aromatic hydrocarbon isomerization processes, the relevant portions of which are incorporated herein by reference in their entirety. Any suitable LPI catalysts known in the art may be used in the processes of this disclosure.
The isomerization catalyst useful in LPI reactors of the processes of this disclosure can comprise a molecular sieve such as a zeolite. Such zeolite can be selected from, but are not limited to, zeolites having the following framework structures, and combinations thereof: MFI, MEL, MWW, MOR, and the like. Preferably, the isomerization catalyst comprises a MFI framework zeolite such as ZSM-5. Preferably, the isomerization catalyst comprises a zeolite having a 10- or 12-member ring structure such as a MWW or MOR framework zeolite, or a mixture thereof. Preferably, the isomerization catalyst comprises both a first zeolite having a MFI framework structure such as ZSM-5, and a second zeolite differing from the first zeolite having a10- or 12-member ring structure. Non-limiting examples of MWW zeolites useful for the isomerization catalyst used in the processes of this disclosure include: MWW-22, MWW-49, MWW-54, and combinations thereof. The zeolites present in the isomerization catalyst may be advantageously in the hydrogen form. The following combinations of zeolites are particularly advantageous for the isomerization catalyst: ZSM-5 and MWW-22; ZSM-5 and MWW-49; and ZSM-5 and MWW-56.
The ZSM-5 zeolite useful for the isomerization catalyst can have one or more of the following characteristics: in the hydrogen form (HZSM-5); having a crystal size≤0.1 micron; having a mesoporous surface area (MSA)≥45 m2/g; a total surface area to mesoporous surface area ratio≤9; and a silica to alumina molar ratio in the range of 20 to 50.
In certain embodiments, the isomerization catalyst can comprise a first metal element selected from Fe, Co, Ni, Ru, Rh, Pd, Re, Os, Ir, Pt, and combinations thereof, and optionally a second metal selected from Sn, Zn, Ag, and combinations thereof. The first metal element may catalyze hydrogenation of olefins that may be produced in the isomerization reactions by, e.g., dealkylation of ethylbenzene. The second metal element may promote the catalytic effect of the first metal element. In other embodiments, the isomerization catalyst is free of precious metal (i.e., Ru, Rh, Pd, Os, Ir, and Pt). In other embodiments, the isomerization catalyst may be free of any Group 7-10 metal. In other embodiments, the isomerization catalyst maybe free of any Group 7-15 metals except aluminum.
The isomerization catalyst may be a self-bound zeolite catalyst substantially free of a binder. Alternatively, the isomerization catalyst can comprise, in addition to the molecular sieves such as zeolites, a binder material. Examples of suitable binder materials include, but are not limited to: alumina, silica, aluminosilicate, zirconia, zircon, titania, clay (e.g., montmorillonite, bentonite, subbentonite), kaolin (e.g., Dixie, McNamee, Georgia and Florida clays or others in which the main mineral constituent is halloysite, kaolinite, nacrite or anauxite), combinations thereof, chemical compounds thereof, and the like. Any suitable quantity of binder may be present in the isomerization catalyst. For example, the binder material may be included in the isomerization catalyst at a concentration from c1 to c2 wt %, based on the total weight of the catalyst, where c1 and c2 can be, independently, 1, 5, 10, 15, 20, 25, 30, 35, 40, 45, 50, 55, 60, 65, 70, 75, 80, 85, 90, 95, 98, 99, as long as c1<c2. The inclusion of a binder in the isomerization catalyst can enhance its mechanical strength, among others.
The isomerization catalyst can be made by processes known in the art. For example, components of the catalyst such as the zeolite and the optional binder can be admixed to form an intimate mixture, which is then extruded to desired shape, dried, calcined, and optionally selectivated to produce the isomerization catalyst. A raw catalyst may be activated in the isomerization reactor or outside of the reactor before the C8 aromatic hydrocarbon isomerization operation. Activation can be conducted by, e.g., exposing the catalyst to a stream of molecular-hydrogen-containing gas.
The isomerization catalyst may be a freshly made catalyst or a regenerated catalyst, or a mixture thereof. Regeneration of the catalyst may be conducted in the isomerization reactor after the catalyst activity has decreased to a threshold level at the end of catalyst cycle by, e.g., exposing the catalyst to a stream of molecular hydrogen-containing gas. Alternatively, ex situ regeneration of the catalyst may be implemented, where the spent catalyst is taken out of the isomerization reactor, heated in an oxygen-rich environment and/or exposed to a molecular hydrogen-containing gas stream to abate coke on its surface.
In some embodiments, the reactor can be a fixed bed reactor, a fluidized bed reactor, or a moving bed reactor. The hydrocarbon liquid in the reactor can flow upward, downward, or in a radial fashion.
The Isomerization Effluents
As a result of the isomerization reactions in the first and/or second isomerization reactor in the isomerization process of this disclosure, the first and/or the second isomerization effluents desirably comprise p-xylene at a concentration higher than in the isomerization hydrocarbon feed, and m-xylene and o-xylene at a combined concentration lower than in the isomerization hydrocarbon feed. The first or the second isomerization effluent where only one of the first and second isomerization reactors is online, or the second isomerization effluent where both the first and second isomerization reactors are online in a lead-lag configuration, may comprise p-xylene at a concentration of from c(pX)1 to c(pX)2 wt %, based on the total weight of the xylenes in the isomerization effluent, where c(pX)1 and c(pX)2 can be, independently, e.g., 15, 16, 17, 18, 19, 20, 21, 22, 23, as long as c(pX)1<c(pX)2. Preferably c(pX)1=18. Preferably c(pX)1=20. Preferably c(pX)1=21. Preferably c(pX)1=22. Preferably c(pX) is in proximity to its concentration in an equilibrium xylenes mixture at the isomerization temperature. In the preferred embodiments of processes of this disclosure where the first and/or second isomerization reactors are LPI reactors and molecular hydrogen is co-fed into one or both of the LPI reactors at a feeding rate≥100 ppm by weight, based on the total weight of the isomerization hydrocarbon feed, one can achieve a high activity of the first and/or second isomerization catalysts, enabling and sustaining c(pX)1≥20 wt % for a long period of time, even at high WHSV of ≥5 hour−1, ≥7.5 hour−1, ≥10 hour−1, ≥12.5 hour−1, ≥15 hour−1, ≥17.5 hour−1, and even ≥20 hour−1, which is totally surprising.
The first and the second isomerization effluents comprise m-xylene and o-xylene at a combined concentration lower than the isomerization hydrocarbon feed. Desirably the m-xylene concentration in the first or second isomerization effluent where only one of the first and second isomerization reactors is online, or in the second isomerization effluent where both the first and second isomerization reactors are online in a lead-lag configuration, based on the total weight of the xylenes in the isomerization effluent, is in proximity to its concentration in an equilibrium xylenes mixture. Thus, the first and/or the second isomerization effluents may comprise m-xylene at a concentration of from c(mX)1 to c(mX)2 wt %, based on the total weight of the xylenes in the isomerization effluent, where c(mX)1 and c(mX)2 can be, independently, e.g., 40, 42, 44, 45, 46, 48, 49, 50, 51, 52, 54, 55, 56, 58, 60, 62, 64, 65, 66, 68, 70, as long as c(mX)1<c(mX)2. Preferably c(mX)2=60. Preferably c(mX)2=58. Preferably c(mX)2=55. Preferably c(mX)2=53. Preferably c(mX)2=50. Desirably the o-xylene concentrations in the first and/or the second isomerization effluents, based on the total weight of the xylenes in the isomerization effluent, are in proximity to its concentration in an equilibrium xylenes mixture. Thus, the isomerization effluent may comprise o-xylene at a concentration of from c(oX)1 to c(oX)2 wt %, based on the total weight of the xylenes in the isomerization effluent, where c(oX)1 and c(oX)2 can be, independently, e.g., 15, 16, 18, 20, 22, 24, 25, 26, 28, 30, 32, 34, 35, as long as c(oX)1<c(oX)2. Preferably c(oX)1=20 and c(oX)2=30. Preferably c(oX)1=20 and c(oX)2=26.
The isomerization effluents can comprise ethylbenzene at various concentrations, depending on the ethylbenzene concentration present in the isomerization hydrocarbon feed.
Recovery of p-Xylene
The isomerization effluent(s), rich in p-xylene and depleted in m-xylene and o-xylene combined compared to the isomerization hydrocarbon feed, can be fed into a p-xylene recovery sub-system, from which a high-purity p-xylene product can be produced.
The p-xylene recovery system can utilize an adsorption chromatography process to separate p-xylene from m-xylene, o-xylene, and ethylbenzene present in the isomerization effluents. Exemplary adsorption chromatography process and systems are described in, e.g., U.S. Pat. Nos. 3,040,777, 3,201,491, 3,422,848, 9,302,201, 3,761,533, 4,029,717, 6,149,874, and 9,302,201, the relevant portions thereof are incorporated herein in their entirety.
The p-xylene recovery system can utilize a crystallization process to separate p-xylene from the other C8 aromatic hydrocarbons present in the isomerization effluents. Exemplary crystallization separation processes and systems are described in, e.g., U.S. Pat. Nos. 3,662,013, 5,329,061, 5,498,822, 6,147,272, and 6,600,083, the relevant portions thereof are incorporated herein in their entirety.
The p-xylene recovery system can utilize a combination of an adsorption chromatography process and a crystallization process as described above.
The p-xylene recovery system can receive, in addition to the isomerization effluent(s) produced from an isomerization reactor in a process of this disclosure, other p-xylene-containing streams, e.g., one or more p-xylene-rich streams produced from a C7−/C9+ aromatic hydrocarbon transalkylation process, a toluene disproportionation process, a benzene/toluene methylation with methanol process, and the like.
In the following cases (Case Nos. 1-7), description is given to various specific embodiments of the processes of this disclosure. The simulation in Cases 1-5 are applicable for LPI processes with cofeeding molecular hydrogen into the LPI reactor pursuant to embodiments of this disclosure, and similarly for LPI processes without cofeeding molecular hydrogen into the isomerization reactor.
Case 1: Manipulating Temperature to Minimize Catalyst Load Size While Optimizing Extent of Xylene Isomerization Versus Undesired Xylenes Loss Reactions
Reactor temperature may be optimized to reduce catalyst load size (increase WHSV) while maintaining low/acceptable xylenes losses.
Continuing with the examples provided in the above section,
The above examples demonstrate the ability to optimize load size and operating temperature to meet a target p-Xylene/Xylenes ratio in the isomerization effluent while reducing capital investment. One skilled in the art will recognize that optimizing in the reverse direction is also an option. One can opt for higher catalyst load and reduce xylenes loss reactions per pass by reducing temperature. However, as shown in
Case 2: Managing Catalyst Activity Loss/Aging with Temperature
Reactor temperature may also be used as a means of maintaining/optimizing reactor yields in response to catalyst activity loss (i.e., aging).
Case 3: Determining Relative Activity Loss for a Given Reactor or Series of Reactors
One can estimate the relative activity loss of a single catalyst system or multiple catalyst systems by monitoring various performance parameters as a function of temperature, feed rate, WHSV, and/or feed concentration which may include but are not limited to p-Xylene/Xylenes ratio in product, delta p-Xylene/Xylenes ratio across reactor, xylenes loss, toluene in product, benzene in product, trimethylbenzenes in product, methylethylbenzenes in product, total C9 aromatics (“A9”) in product, total C10 aromatics (“A10”) in product, total C9+ aromatics (“A9+”) in the product, delta toluene across reactor, delta benzene across reactor, delta trimethylbenzenes across reactor, delta methylethylbenzenes across reactor, delta total A9 across reactor, delta total A10 across reactor, delta total A9+ across reactor, ethylbenzene conversion across reactor, non-aromatics conversion across the reactor (individual species, subset of species, or total non-aromatics), C5− products (individual species, subset of species, or total C5−), delta C5− across reactor (individual species, subset of species, or total C5−), or any combination of these parameters.
As an example,
In the example of a lead-lag or multiple reactor system, the estimated relative activity loss or any of the parameters used in determining the estimated activity loss may be used to indicate/signal when to take a reactor/catalyst bed out of service. In the example provided in
Case 4: Catalyst Activity De-edging at Start of Run; Catalyst Activity Management During Operation Rate Turndown
Opposite to Case 2, at start of cycle for a fresh catalyst bed it is common to have higher than desired initial catalyst activity that results in higher xylenes loss or undersirable levels of byproducts such as benzene, toluene, or A9+ molecules. This can also be the case at middle or end of run, if unit rates are substantially turned down below the design WHSV. To minimize the make of undesired byproducts at start of cycle or during significant turndown operation, the temperature can be lowered to reduce the undesired byproduct yield. In the case of “de-edging” initial fresh catalyst activity, the temperature can be gradually increased as the byproduct production rate declines over typically the first 1-2 months, but possibly up to 1 or more years for extreme cases. In the case, of unit rate turndown, the temperature can be adjusted as needed to manage yields. The absolute magnitude of temperature change will be dependent on percentage turndown from design rates as well as the relative activity loss that the catalyst has incurred on stream leading up to the turndown event.
Temperature control to manage higher than desired initial catalyst activity can also be used in multi-bed/multi reactor systems such as a lead-lag configurations. In the example of the lead-lag configuration, the fresh catalyst bed (whether lead or lag) can be started up at lower temperature versus the reactor that has already been on stream. The independent temperature control to the individual reactors can be managed by any number of configurations which may include but is not limited to; independent heaters/coolers on feed to each reactor, independent bypass lines around heaters/coolers to each reactor allowing for temperature control to each reactor, injection of a cooled or heated fresh stream to the lag reactor/reactors that is mixed with the effluent of the preceding reactor/reactors to control temperature to the lag reactor/reactors, utilizing natural heat loss between the lead-lag reactor/reactors to operate the lag reactor/reactors cooler than the lead reactor/reactors. Any combination of these temperature control methods may be used.
Case 5: Managing Lead Lag Configurations Where the Reactor/Catalyst Bed Size Varies Between Lead and Lag
As shown in Case 1, the reactor operating temperature can be manipulated to allow similar yields for reactors of varying catalyst load or WHSV. In a lead-lag or multi-reactor configuration, the reactors may be optimized/designed to have different quantities of catalyst. See TABLE II, for Cases 5a and 5b which are provided for illustration. In the TABLEs of this disclosure, for the purpose of brevity, EB means ethylbenzene, L-Paraffins means linear paraffins, BCNA means branched or cyclic non-aromatic hydrocarbons, MEBZ means methylethylbenzenes, DEBZ means diethylbenzenes, TriMBZ means trimethylbenzenes, and DMEBZ means dimethylethylbenzenes.
Case 5a estimates the yields for a lead-lag configuration where the lead catalyst bed operates at 5 hour−1 WHSV and 240° C. and the lag catalyst bed operates at 10 hour−1 WHSV and 260° C. In this Case, the bulk of the isomerization reaction occurs in the larger lead catalyst bed which allows for the use of a smaller lag bed operating at 10 hour−1 WHSV and a slightly elevated temperature of 260° C. to serve as a finishing reactor for the xylene isomerization. Please note this combined lead-lag system with varied WHSV and temperature yields similar results to the larger single bed design case of 2.5 hour−1 WHSV and 240° C. that is presented in the data in
Case 5b estimates the yields for a lead-lag configuration where the lead and lag catalyst beds operate at equivalent 10 hour−1 WHSV each, but the operating temperatures are manipulated between the lead and lag reactors (240° C. for lead reactor, 270° C. for lag reactor). This configuration results in similar p-Xylene/Xylenes ratio in the product as Case 5a, but does have marginally higher estimated xylenes losses of 0.53% vs 0.46% xylenes loss for Case 5a.
Case 6: Managing Hydrogen Co-Feed to Lead-Lag or Multi Reactor Systems
Low concentration hydrogen injection has been surprisingly found to attenuate/mitigate activity loss in liquid phase isomerization systems. It should be noted that the level of hydrogen co-feed to a lead-lag or multi reactor system can be controlled for only the inlet of the first reactor, or it can be separately controlled to each individual reactor or any subset of reactors/catalyst beds. The hydrogen level in the feed to each reactor may be controlled by any of the following means, but is not limited to the subset listed herein; independent hydrogen injection to any given reactor/reactors, addition of fresh feed to subsequent reactor/reactors that is combined with at least a portion of the effluent from the previous reactor/reactors, flashing of the effluent of a previous reactor/reactors to a lower pressure to remove hydrogen from the liquid phase prior to feeding at least a portion of the remaining liquid stream to the subsequent reactor/reactors.
Hydrogen co-feed may be optimized to individual reactors or any a subset of reactors to minimize aging at varying WHSV, Temperature, or Feed Concentrations. It may also be manipulated/optimized over the cycle length of the catalyst in response to activity decline. For example, at start of cycle, the hydrogen co-feed may be reduced if the fresh catalyst activity to a given bed is undesirably high and resulting in high undesired byproduct make. Reducing the hydrogen co-feed during this “de-edging” period will help to promote/accelerate initial activity loss to reduce the period of high undesired byproduct make at start of cycle. As the undesired byproduct make declines, the hydrogen co-feed may be increased to reduce long term activity loss and to ultimately extend the run length of catalyst bed.
Hydrogen co-feed may be optimized through the cycle if increased rate of activity loss is observed at various operating conditions. For example, during periods where the observed rate of activity decline is high, the hydrogen co-feed may be increased. In periods where the observed rate of activity decline is low/negligible, the hydrogen co-feed may be decreased to reduce operating costs.
Case 7: Optimizing WHSV, Temperature, Hydrogen Co-Feed for Single Bed or for Lead-Lag or Multi Reactor Systems Based on Feed Composition
Feed concentration impacts both p-Xylene/Xylenes ratio in the isomerization effluent and the extent of undesired byproduct make. The concepts outlined in cases 1-6 above can also be applied to optimize yield and capital where significant feed compositional differences or variance is expected. Please see TABLE III for an illustration where a single bed reactor temperature is manipulated to achieve similar yields for a standard feed composition versus a high ortho-xylene containing feed. In this example, the standard feed contains low p-xylene (<1 wt %) content, low to moderate EB content (˜5.3 wt %) and moderate o-xylene content (˜37 wt %). The Ox-rich feed also contains low p-xylene (<1%), but also has low EB and m-xylene content (<1 wt % for both). In this example, when processing the Ox-rich feed, the single catalyst bed yields lower p-Xylene/Xylenes ratio in the product and lower xylenes loss when compared to processing standard feed at the same conditions of 2.5 hour−1 WHSV and 240° C. However, in this example, the p-Xylene/Xylenes ratio in the product for the ortho-xylene rich feed case can be raised to match the standard feed product by raising the reactor temperature from 240° C. to 250° C. As discussed in the above sections, optimization of WHSV, temperature, or a combination of both can be used to achieve equivalent product to the standard feed case.
The isomerization processes of this disclosure can be advantageously used in the process illustrated in
As shown in
This disclosure is further illustrated by the following non-limiting examples.
EXAMPLESExamples 1 to 3 demonstrate the advantages of a specific LPI process comprising co-feeding molecular hydrogen at a feeding rate≥100 ppm by weight, based on the total weight of the isomerization hydrocarbon feed, in comparison to LPI processes without co-feeding molecular hydrogen, and LPI processes comprising co-feeding molecular hydrogen at a feeding rate≤10 ppm by weight. In Examples 1 to 3, an isomerization hydrocarbon feed consisting essentially of ethylbenzene, p-xylene, o-xylene and ethylbenzene was fed into a LPI reactor pre-loaded with a given quantity of a ZSM-5-based isomerization catalyst. A molecular hydrogen cylinder was connected to the LPI reactor or the isomerization hydrocarbon feed source to enable co-feeding of it to the reactor at various feeding rates. The isomerization hydrocarbon feed was heated to have a reactor inlet temperature in the range from 200 to 300° C. The pressure in the reactor was in the range from 689 to 3447 kPa (gauge). Under such conditions, C8 aromatic hydrocarbon present in the reactor was substantially entirely in liquid phase, and molecular hydrogen co-fed into the reactor, if any, was substantially entirely dissolved in the liquid phase of the hydrocarbons. The isomerization conditions were varied to test the deactivation of the isomerization catalyst. The isomerization effluent exiting the isomerization reactor was then analyzed by using gas chromatography to determine the concentrations (wt %) of p-xylene, o-xylene, and m-xylene, based on the total weight of the isomerization effluent. The measured p-xylene concentration (C(pX), wt %), o-xylene concentration (C(oX), wt %), and m-xylene concentration (C(mX), wt %) were then used to calculate the p-xylene concentration based on the total quantity of xylenes (C(pX/X), %) according to the following formula:
In the description of Examples 1 to 3, “CGpGC” means cumulative weight of the isomerization hydrocarbon feed in grams per gram of the isomerization catalyst that has been processed by the batch of the isomerization catalyst. In a hypothetical case where the feeding rate of the isomerization hydrocarbon feed is maintained constant, CGpGC would correspond to the product of the feeding rate and time on stream of the catalyst.
Example 1In this example, the isomerization hydrocarbon feed comprised 3.4 wt % of p-xylene, 64.1 wt % of m-xylene, 18.3 wt % of o-xylene, 12.3 wt % of ethylbenzene, and 1.3 wt % of other hydrocarbons. In the isomerization process, isomerization conditions such as WHSV and feeding rate of molecular hydrogen were varied. The reactor pressure was varied in order to maintain the molecular hydrogen co-fed at various feeding rate into the isomerization reactor substantially entirely dissolved in the liquid phase C8 aromatic hydrocarbons. The reaction temperature in the isomerization reactor was maintained at approximately 280° C. p-Xylene concentration in the isomerization effluent as a function of CGpGC in the process is shown in
As can be seen in
At the end of CGpGC range 203 and the beginning of range 205, the molecular hydrogen feeding rate was increased to 255 ppm while the WHSV was maintained at 10 hour−1, which were then maintained throughout range 205. At the beginning of range 205, the C(pX) increased substantially compared to at any CGpGC in ranges 203 and 201, indicating a substantial increase of the activity of the isomerization catalyst in the presence of a substantially increased molecular hydrogen concentration in the reactor. Throughout range 205, the C(pX) was substantially steady, indicating little deactivation of the isomerization catalyst in the presence of molecular hydrogen at 255 ppm at a high WHSV of 10 hour−1.
At the end of CGpGC range 205, the WHSV was ramped up while the feeding rate of molecular hydrogen was maintained at 255 ppm. During range 207, the WHSV and the feeding rate of molecular hydrogen were maintained at 20 hour−1 and 255 ppm, respectively. The C(pX) in range 207 was lower compared to in range 205, which is understandable due to the twice as high of WHSV in range 207. Nonetheless, the C(pX) throughout range 207 was significantly higher than at the ends of ranges 201 and 203, even though the WHSV in range 207 is twice as high as in ranges 201 and 203, clearly demonstrating a significantly higher activity of the isomerization catalyst in range 207 than in ranges 201 and 203, and the effect of a high molecular hydrogen concentration in the isomerization reactor on the activity of the catalyst. Moreover, during range 207, once the C(pX) stabilized, it decreased very little, showing a very low deactivation rate of the isomerization catalyst even at a very high WHSV of 20 hour−1, indicating the substantial effect imparted by the high concentration of molecular hydrogen in the isomerization reactor.
At the end of range 207, the feeding rate of molecular hydrogen was reduced. During range 209, no molecular hydrogen was fed into the reactor, and the WHSV was maintained a 20 hour−1. As a result of the absence of molecular hydrogen in the isomerization reactor, the C(pX) decreased substantially in range 209 compared to in range 207, again demonstrating the effect of the presence of molecular hydrogen in the isomerization reactor on the activity of the isomerization catalyst. The C(pX) in range 209 was similar to at the end of range 203 where no molecular hydrogen was fed into the reactor as well.
At the end of range 209, a catalyst deactivating agent was injected into the reactor to expedite the deactivation of the isomerization catalyst. Afterwards, during range 211, the WHSV was maintained at 20 hour−1 while no molecular hydrogen was co-fed into the reactor. C(pX) in range 211 was low as a result of the deactivated catalyst, a high WHSV of 20 hour−1, and the absence of molecular hydrogen co-fed into the reactor.
This Example 1 clearly demonstrates that in a liquid-phase LPI process, by co-feeding molecular hydrogen at a feeding rate of ≥100 ppm, based on the total weight of the C8 aromatic hydrocarbon isomerization feed, the activity of the isomerization catalyst was enhanced, and the deactivation of the catalyst was reduced substantially, compared to feeding molecular hydrogen at a low feeding rate or no co-feeding of molecular hydrogen. Given the high activity and exceedingly low deactivation rate of the isomerization catalyst shown at high molecular hydrogen concentration, a LPI process at very high WHSV of, e.g., ≥10 hour−1, ≥15 hour−1, ≥17.5 hour−1, and even ≥20 hour−1, can be enabled by co-feeding molecular hydrogen into the isomerization reactor at ≥100 ppm.
Example 2 ComparativeIn this example, the isomerization hydrocarbon feed comprised 1.1 wt % of p-xylene, 67.7 wt % of m-xylene, 29.8 wt % of o-xylene, 1.0 wt % of ethylbenzene, and 0.5 wt % of other hydrocarbons. In the isomerization process, the WHSV was varied, the feeding rate of molecular hydrogen was maintained at 9 ppm by weight, based on the weight of the isomerization hydrocarbon feed. The reaction temperature in the isomerization reactor was maintained at approximately 239° C. The reactor pressure was sufficient to maintain the molecular hydrogen co-fed into the isomerization reactor substantially entirely dissolved in the liquid phase C8 aromatic hydrocarbons. p-Xylene concentration in the isomerization effluent based on the total quantity of xylenes (C(pX/X), wt %) as a function of WHSV (hour−1) in the process is shown in
The isomerization hydrocarbon feed used in this example is deemed as relatively easy to isomerize due to the low concentration of ethylbenzene therein. Thus, one would appreciate similar performance of the catalyst at high WHSV. Nonetheless, as
In this example, the isomerization hydrocarbon feed comprised 8.1 wt % of p-xylene, 58.1 wt % of m-xylene, 24.4 wt % of o-xylene, 7.5 wt % of ethylbenzene, and 1.9 wt % of other hydrocarbons. In the isomerization process, the WHSV was maintained at 4 hour−1. No molecular hydrogen was fed into the isomerization reactor. The reaction temperature in the isomerization reactor was maintained at approximately 239° C. p-Xylene concentration in the isomerization effluent based on the total quantity of xylenes (C(pX/X), wt %) as a function of Time on Stream (TOS, day) in the process is shown in
The feed used in this example is deemed as relatively easy to isomerize due to the high concentration of p-xylene therein. Thus, one would expect a high performance of the catalyst at a WHSV of 4 hour−1. Nonetheless, as
With this example taken into consideration, the positive effect of co-feeding molecular hydrogen at a high feeding rate of ≥100 ppm in Example 1 in reducing deactivation of the isomerization catalyst is further demonstrated.
Example 4 InventiveIn this prophetic example, an overall p-xylene production process in an aromatics production complex including two LPI reactors operating with WHSV in the range from 10 to 20 hour−1 in series pursuant to the process of this disclosure is simulated. An isomerization hydrocarbon feed comprising C8 aromatic hydrocarbons is fed to a first LPI reactor equipped with a first isomerization catalyst operated under a first set of LPI conditions to produce a first isomerization effluent. A second LPI reactor equipped with a second isomerization catalyst substantially similar to the first isomerization catalyst is initially bypassed. For a first period of time, the first isomerization effluent (or a portion thereof) is supplied to a xylenes splitter and/or a p-xylene recovery sub-system directly. When the activity of the first isomerization catalyst decreases to a threshold level at the end of the first period of time, the second LPI reactor was placed online as a lag reactor, i.e., the first isomerization effluent is fed into the second LPI reactor operated under a second set of isomerization conditions to produce a second isomerization effluent. The second isomerization effluent (or a portion thereof) is supplied to the xylenes splitter and/or the p-xylene recovery sub-system directly. The serial operation of the two LPI reactors allows the operation of the first isomerization catalyst to be extended. After operation of the first and second reactors in lead-lag configuration for a second period of time, while the second isomerization catalyst is still in the early stage or the middle of its cycle and the first isomerization catalyst has reached the end of its cycle life to become a spent first isomerization catalyst, the isomerization hydrocarbon feed is directly fed into the second LPI reactor bypassing first LPI reactor is bypassed. The second effluent produced from the second LPI reactor (or a portion thereof) is fed to the xylenes splitter and/or the p-xylene recovery sub-system directly. The spent first isomerization catalyst can be regenerated in situ in the first LPI reactor, or taken out for ex-situ regeneration. A fresh and/or regenerated isomerization catalyst can be installed in the offline first LPI reactor while the second LPI reactor continues its isomerization operation for a third period of time. When the activity of the second isomerization catalyst in the second LPI has decreased to a threshold level at the end of the third period of time, and the first LPI is already equipped with a regenerated and/or replaced batch of catalyst ready for isomerization operation, the first LPI reactor can be placed online as a lag reactor downstream of the second LPI reactor as a lead reactor. Thus, the second isomerization effluent produced from the second, lead LPI reactor is fed into the first, lag LPI reactor, and the first isomerization effluent produced from the lag LPI reactor, or a portion thereof, can be fed to a xylenes splitter and/or a p-xylene recovery sub-system directly. The above process can be repeated to allow for continuous isomerization operation in the aromatics production complex until a major site turnaround interval is achieved. In contrast, in a comparative process utilizing a single LPI reactor in the xylenes loop, each shutdown of the LPI reactor to regenerate or replace the spent isomerization catalyst can cause 2-3 weeks of complete shutdown of the isomerization process, resulting in significant p-xylene productivity loss if the rest of the aromatics production complex continues, or the shutdown of the aromatics production complex if such productivity loss is to be avoided.
Compared to the comparative process using a single LPI reactor, the process of this disclosure using two LPI reactors arranged in changeable lead-lag configuration have many additional significant advantages: (i) because the two LPI reactors in the process of this disclosure can each have a design capacity substantially smaller than the single LPI reactor in the comparative process handling the same quantity of isomerization hydrocarbon feed, the total capital cost for the two LPI reactors in the process of this disclosure can be ≤87% of the capital cost for the single LPI reactor in the comparative process; (ii) the process of this disclosure reduces required catalyst inventory by 50% to 75% compared to the comparative process; (iii) the total quantity of spent catalyst waste in the process of this disclosure is smaller by regenerating and reusing the catalyst over multiple cycles; (iv) the process of this disclosure improves theoretical utilization of equipment over major turnaround interval by ≥2%; and (v) the process of this disclosure eliminates the need for oversizing/overdesigning parallel vapor isomerization facilities to manage liquid phase isomerization unit outages which reduces capital investment and catalyst costs and improves the design yields of the vapor isomerization catalyst system.
LPI reactors operated with a high WHSV, e.g., a WHSV≥5, 7.5, 10, 12.5, 15, 17.5, or even 20 hour−1, such as those enabled by co-feeding molecular hydrogen at a feeding rate of ≥100 ppm by weight, based on the total weight of the isomerization hydrocarbon feed, as illustrated in Examples 1 to 3 above, can be advantageously used in the isomerization processes of this disclosure where two or more LPI reactors are arranged in changeable lead-lag configurations. The additional benefit of reduced deactivation rate of the isomerization catalyst imparted by co-feeding a substantial quantity of molecular hydrogen into the LPI reactor, even at high WHSV, can further significantly improve the capital and operation costs of the process of this disclosure.
Claims
1. An isomerization process, the process comprising:
- (I) providing an isomerization hydrocarbon feed comprising C8 aromatic hydrocarbons;
- (II) providing a first isomerization reactor comprising a first isomerization catalyst disposed therein, and a second isomerization reactor comprising a second isomerization catalyst disposed therein;
- (III) feeding the isomerization hydrocarbon feed and optionally molecular hydrogen (H2) into the first isomerization reactor;
- (IV) contacting the isomerization hydrocarbon feed and the optional molecular hydrogen with the first isomerization catalyst under a first set of isomerization conditions to produce a first isomerization effluent exiting the first isomerization reactor for a first period of time, the first period of time being shorter than the total life cycle of the first isomerization catalyst;
- (V) obtaining a p-xylene product stream from at least a portion of the first isomerization effluent during the first period of time;
- (VI) at the end of the first period of time, feeding at least a portion of the first isomerization effluent and optionally additional molecular hydrogen into the second isomerization reactor, wherein at the end of the first period of time, the second isomerization catalyst has a prospective runtime longer than the first isomerization catalyst;
- (VII) after step (VI), contacting at least a portion the first isomerization effluent with the second isomerization catalyst under a second set of isomerization conditions in the second isomerization reactor to produce a second isomerization effluent exiting the second isomerization reactor for a second period of time;
- (VIII) continuing step (IV) during the second period of time;
- (IX) obtaining a p-xylene product stream from at least a portion of the second isomerization effluent during the second period of time;
- (X) at the end of the life cycle of the first isomerization catalyst, where the first isomerization catalyst becomes a spent first isomerization catalyst, feeding the isomerization hydrocarbon feed and optionally molecular hydrogen into the second isomerization reactor, and stopping feeding the isomerization hydrocarbon feed into the first isomerization reactor;
- (XI) after step (X), contacting the isomerization hydrocarbon feed with the second isomerization catalyst under the second set of isomerization conditions to produce a third isomerization effluent exiting the second isomerization reactor for a third period of time; and
- (XII) obtaining a p-xylene product stream from at least a portion of the third isomerization effluent during the third period of time.
2. The isomerization process of claim 1, further comprising:
- (XIII) during the third period of time, regenerating the spent first isomerization catalyst in the first isomerization reactor, and/or replacing at least a portion of the spent first isomerization catalyst in the first isomerization reactor with a fresh batch of the first catalyst and/or a regenerated batch of the first catalyst.
3. The isomerization process of claim 1, wherein the third period of time ends before the end of the life cycle of the second isomerization catalyst.
4. The isomerization process of claim 3, further comprising:
- (XIV) after step (XIII), designating the second isomerization reactor as the first isomerization reactor, the second isomerization catalyst as the first isomerization catalyst, the second set of isomerization conditions as the first set of isomerization conditions, the first isomerization reactor as the second isomerization reactor, the first isomerization catalyst as the second isomerization catalyst, the first set of isomerization conditions as the second set of isomerization conditions, and subsequently repeating steps (III), (IV), (V), (VI), (VII), (VIII), (IX), (X), (XI), and (XII).
5. The isomerization process of claim 4, wherein step (XIV) further comprises repeating step (XIII).
6. The isomerization process of claim 5, further comprising:
- (XV) stopping operating the first isomerization reactor and/or the second isomerization reactor at a major turn-around point of time of the respective isomerization reactor(s).
7. The isomerization process of claim 1, wherein the first period of time is at least half of the life cycle of the first isomerization catalyst.
8. The isomerization process of claim 1, wherein at the end of the first period of time, in step (IV), the activity of the first isomerization catalyst is no greater than 50% of the average activity of the first isomerization catalyst during the first period of time.
9. The isomerization process of claim 1, wherein:
- the first set of isomerization conditions are such that vapor-phase isomerization is carried out in the first isomerization reactor; and
- the second set of isomerization conditions are such that vapor-phase isomerization is carried out in the second isomerization reactor.
10. The isomerization process of claim 1, wherein:
- the first set of isomerization conditions are such that liquid-phase isomerization is carried out in the first isomerization reactor; and
- the second set of isomerization conditions are such that vapor-phase isomerization is carried out in the second isomerization reactor.
11. The isomerization process of claim 1, wherein:
- the first set of isomerization conditions are such that liquid-phase isomerization is carried out in the first isomerization reactor; and
- the second set of isomerization conditions are such that liquid-phase isomerization is carried out in the second isomerization reactor.
12. The isomerization process of claim 1, wherein the first isomerization catalyst and the second isomerization catalyst, when fresh, have substantially the same composition.
13. The isomerization process of claim 11, wherein liquid-phase isomerization is carried out in both the first and second isomerization reactor, and the first and second isomerization conditions comprise at least one of the following:
- a temperature in a range from 200 to 300° C.;
- a weight hourly space velocity in a range from 1.5 to 20 hour−1; and
- a gauge pressure in a range from 0 to 3500 kilopascal.
14. The isomerization process of claim 13, wherein the first and second isomerization conditions comprise at least one of the following:
- a temperature in a range from 240 to 300° C.; and
- a weight hourly space velocity in a range from 2.5 to 15 hour−1; and
- a gauge pressure in a range from 690 to 3500 kilopascal.
15. The isomerization process claim 11, wherein:
- in step (III), molecular hydrogen is fed into the first isomerization reactor at a first hydrogen feeding rate; and/or
- in step (X), molecular hydrogen is fed into the second isomerization reactor at a second hydrogen feeding rate.
16. The isomerization process of claim 15, wherein the first hydrogen feeding rate and the second hydrogen feeding rate are in a range below 100 ppm by weight, based on the total weight of the isomerization hydrocarbon feed.
17. The isomerization process of claim 16, wherein the first set of isomerization conditions and/or the second set of isomerization conditions comprise a WHSV of less than 5 hour−1.
18. The isomerization process of claim 15, wherein the first hydrogen feeding rate and the second hydrogen feeding rate are in a range from 100 to 5000 by weight, based on the total weight of the isomerization hydrocarbon feed.
19. The isomerization process of claim 18, wherein the first set of isomerization conditions and/or the second set of isomerization conditions comprise a WHSV from 5 to 25 hour−1.
20. The isomerization process of claim 18, wherein the first set of isomerization conditions comprises a first WHSV, and the second set of isomerization conditions comprise a second WHSV, and the first WHSV is lower than the second WHSV.
21. The isomerization process of claim 11, wherein the first set of isomerization conditions comprises a first temperature, and the second set of isomerization conditions comprises a second temperature, and the first temperature is lower than the second temperature.
22. The isomerization process of claim 13, wherein additional molecular hydrogen is not fed into the second isomerization reactor in step (VI).
23. A C8 aromatic hydrocarbon isomerization process, the process comprising:
- (I) providing an isomerization hydrocarbon feed comprising C8 aromatic hydrocarbons;
- (II) providing a first isomerization reactor comprising a first isomerization catalyst disposed therein, and a second isomerization reactor comprising a second isomerization catalyst disposed therein;
- (III) feeding the isomerization hydrocarbon feed and optionally molecular hydrogen (H2) into the first isomerization reactor;
- (IV) contacting the isomerization hydrocarbon feed and the optional molecular hydrogen with the first isomerization catalyst under a first set of isomerization conditions to produce a first isomerization effluent exiting the first isomerization reactor for a first period of time, the first period of time being shorter than the total life cycle of the first isomerization catalyst;
- (V) obtaining a p-xylene product stream from at least a portion of the first isomerization effluent during the first period of time;
- (VI) at the end of the first period of time, feeding at least a portion of the first isomerization effluent and optionally additional molecular hydrogen into the second isomerization reactor, wherein at the end of the first period of time, the second isomerization catalyst has an prospective runtime longer than the first isomerization catalyst;
- (VII) after step (VI), contacting at least a portion of the first isomerization effluent with the second isomerization catalyst under a second set of isomerization conditions in the second isomerization reactor to produce a second isomerization effluent exiting the second isomerization reactor for a second period of time;
- (VIII) continuing step (IV) during the second period of time;
- (IX) obtaining a p-xylene product stream from at least a portion of the second isomerization effluent during the second period of time;
- (X) at the end of the life cycle of the first isomerization catalyst, where the first isomerization catalyst becomes a spent first isomerization catalyst, feeding the isomerization hydrocarbon feed and optionally molecular hydrogen into the second isomerization reactor, and stopping feeding the isomerization hydrocarbon feed into the first isomerization reactor;
- (XI) after step (X), contacting the isomerization hydrocarbon feed with the second isomerization catalyst under the second set of isomerization conditions to produce a third isomerization effluent exiting the second isomerization reactor for a third period of time;
- (XII) obtaining a p-xylene product stream from at least a portion of the third isomerization effluent during the third period of time;
- (XIII) during the third period of time, regenerating the spent first isomerization catalyst in the first isomerization reactor, and/or replacing at least a portion of the spent first isomerization catalyst in the first isomerization reactor with a fresh batch of the first catalyst and/or a regenerated batch of the first catalyst; and
- (XIV) after step (XIII), designating the second isomerization reactor as the first isomerization reactor, the second isomerization catalyst as the first isomerization catalyst, the second set of isomerization conditions as the first set of isomerization conditions, the first isomerization reactor as the second isomerization reactor, the first isomerization catalyst as the second isomerization catalyst, the first set of isomerization conditions as the second set of isomerization conditions, repeating steps (III), (IV), (V), (VI), (VII), (VIII), (IX), (X), (XI), and (XII), and optionally repeating step (XIII).
24. The isomerization process of claim 23, wherein the first and second isomerization conditions comprise at least one of the following:
- a temperature in a range from 240 to 300° C.;
- a weight hourly space velocity in a range from 2.5 to 15 hour−1; and
- a gauge pressure in a range from 690 to 3500 kilopascal.
25. The isomerization process of claim 23, wherein:
- in step (III), molecular hydrogen is fed into the first isomerization reactor at a first hydrogen feeding rate; and
- in step (X), molecular hydrogen is fed into the second isomerization reactor at a second hydrogen feeding rate.
26. The isomerization process of claim 25, wherein the first hydrogen feeding rate and the second hydrogen feeding rate are in a range from 100 to 5000 by weight, based on the total weight of the isomerization hydrocarbon feed, and the first set of isomerization conditions and the second set of isomerization conditions comprise a WHSV from 5 to 25 hour−1.
Type: Application
Filed: Aug 11, 2020
Publication Date: Sep 15, 2022
Inventors: Robert G. Tinger (Friendswood, TX), Hari Nair (Spring, TX), Todd E. Detjen (Bellaire, TX), Paul Podsiadlo (Humble, TX), Travis D. Sparks (Deer Park, TX)
Application Number: 17/632,584
